CN1874980A - Process for the production of 1,3-propanediol by catalytic hydrogenation of 3-hydroxypropanal in the presence of a hydration co-catalyst - Google Patents

Abstract

A process for preparing 1,3-propanediol comprising the steps of: (a) preparing an aqueous mixture of 3-hydroxypropanal, (b) passing said aqueous 3-hydroxypropanal mixture to a hydrogenation zone wherein the hydrogenation zone is comprised of at least two stages wherein the hydrogenation in the first stage is carried out under hydrogenation conditions at a temperature in the range of 50 to 130 1/2 C in the presence of a fixed bed or slurry hydrogenation catalyst and wherein an acidic co-catalyst is added or present in at least one of the later stages and the hydrogenation in said later stages is carried out under hydrogenation conditions at a higher temperature than that of the first hydrogenation stage and in the range of 70 to 155 1/2 C to form an aqueous solution of 1,3-propanediol; and (c) recovering said 1,3-propanediol.

Description

Process for the preparation of 1,3-propanediol by catalytic hydrogenation of 3-hydroxypropionaldehyde in the presence of a hydration promoter

technical field

The present invention relates to a process for the hydrogenation of 3-hydroxypropanal to 1,3-propanediol. More specifically, the present invention relates to the introduction of a binary catalyst system that promotes hydration and hydrogenation of by-products to 1,3-propanediol to increase process yield.

Background technique

1,3-Propanediol (PDO) can be prepared by hydrogenating an aqueous solution of 3-hydroxypropionaldehyde. 3-Hydroxypropanal (HPA) intermediate solutions can be prepared by a process comprising the steps of catalytic hydroformylation of oxirane (reaction with synthesis gas _H2 /CO) in an organic solvent to form a dilute solution of HPA The mixture, followed by extraction of HPA in water to form a more concentrated HPA solution, and subsequent hydrogenation of HPA to PDO.

US 5786524 (incorporated herein by reference) describes a process wherein ethylene oxide and synthesis gas are contacted at 50-100°C and 500-5000 psig in the presence of a cobalt or rhodium catalyst and catalyst promoter to produce Intermediate product mixture of HPA. Water is added to the HPA mixture, and most of the HPA is extracted into the water to provide an aqueous phase containing a higher concentration of HPA and an organic phase containing at least a portion of the catalyst. The aqueous and organic phases were separated, and the aqueous phase was diluted with an aqueous solution of 1,3-propanediol. This aqueous solution is then passed into a hydrogenation zone to contact a fixed bed hydrogenation catalyst to form a PDO-containing aqueous solution, which is then recovered.

Alternatively, an aqueous intermediate stream of 3-hydroxypropanal can be prepared by hydration of acrolein as described in US 5015789 (incorporated herein by reference).

US 5945570 (incorporated herein by reference) describes a novel hydrogenation catalyst for the hydrogenation of 3-hydroxypropanal. The catalyst is a promoted nickel catalyst which also includes molybdenum and a binder material which may be a silicate or an oxide of silicon, zirconium, aluminum or the like. The catalyst system is characterized by an extended catalyst lifetime in the hydrogenation reaction environment because, relative to existing catalyst systems, such as the supported hydrogenation catalyst system described in US 5,786,524 discussed above, the catalyst The pulverization strength of the system increases.

During the hydrogenation process or in the steps preceding the hydrogenation process, there may be a relatively large concentration of HPA and PDO at the same time. These components can react to form the thermodynamically favorable cyclic acetal MW132:

Further reaction of the cyclic acetal may occur to produce additional by-product species.

It is extremely beneficial if the reverse reaction, that is, the reaction of acetal into PDO, can be promoted during the hydrogenation process, because this will increase the yield of PDO. However, the life-extending catalyst described in US 5945570 does not effectively promote the reverse reaction of acetals back to PDO and HPA (where the latter is then hydrogenated to PDO). The present invention allows efficient reversal of acetal by-products to PDO at improved space-time yields.

Contents of the invention

The present invention is a process for the preparation of 1,3-propanediol comprising the steps of (a) preparing an aqueous mixture of 3-hydroxypropanal and reaction by-products, (b) passing the aqueous 3-hydroxypropanal mixture by adding A hydrogenation zone, wherein said hydrogenation zone consists of at least two stages, wherein the first stage is carried out under hydrogenation conditions at a temperature in the range of 50-130°C in the presence of a fixed bed or slurry hydrogenation catalyst Hydrogenation, and wherein in at least one subsequent stage an acidic cocatalyst is added or present, said subsequent stage is carried out under hydrogenation conditions at a temperature higher than that of the first hydrogenation stage and in the range of 70-155°C hydrogen to form an aqueous solution of 1,3-propanediol, wherein the acidic promoter is selected from the group consisting of acidic zeolites, acidic cation exchange resins, acidic or amphoteric metal oxides, heteropolyacids, and soluble acids; and (c) recovering the 1,3-propanediol. In a preferred embodiment, the hydrogenation zone consists of at least three stages, wherein the second stage of hydrogenation is carried out at a temperature higher than that of the first stage in the range of 70-155° C. (preferably 70-140° C.). Hydrogen, and wherein an acidic cocatalyst is added or present in the last stage of the hydrogenation zone, wherein at a temperature higher than the temperature of the second stage and at a temperature greater than or equal to 120°C (preferably 120°C to 155°C) The final stage of hydrogenation is carried out under hydrogen conditions. In another preferred embodiment, a slurry hydrogenation catalyst and an acidic cation exchange resin are used, and the temperature in the final hydrogenation stage is 100-140°C.

Detailed ways

A hydroformylation process useful for the preparation of aqueous mixtures of 3-hydroxypropionaldehyde from ethylene oxide and synthesis gas is described in detail in US 5786524 (herein incorporated by reference). US5015789 describes similar aqueous 3-hydroxypropanal mixtures prepared by hydration of acrolein.

The aqueous stream input to the hydrogenation section is an aqueous solution containing HPA, based on the weight of the aqueous liquid (usually water), the concentration of HPA is in the range of 3-50 wt%, preferably 10-40 wt%. If a fixed bed catalyst is used for the hydrogenation reaction, it is desirable to dilute the HPA concentration in the feed to the first hydrogenation reactor to a value of 0.2-15 wt%, and most preferably 0.5-8 wt%. While the HPA solution can be diluted to the desired concentration with any aqueous liquid that does not interfere with HPA hydrogenation, including water, it is preferred to use an aqueous solution containing PDO, such as a portion of the product stream from the hydrogenation step. Dilution with this PDO-containing solution serves to concentrate the PDO in the water of the system, thereby avoiding high costs and recovering diluted PDO from the water that may have come from the water used alone as diluent. A preferred dilute stream contains PDO and HPA in an amount in the range of 20-40 wt%. The dilution stream is preferably cooled prior to mixing with the output stream of the hydroformylation to bring the temperature of the combined stream to that required for the initial stage of input to the fixed bed hydrogenation.

If a slurry catalyst is used in a back-mixed reactor, the extent of the first stage reaction can be controlled to obtain a similar concentration range of 0.2-15 wt% HPA, preferably 0.5-8 wt%, thereby eliminating the need for pre-diluted feedstock.

The hydrogenation process can be carried out in one stage or in two or more sequential temperature stages. In a preferred embodiment, the hydrogenation is carried out at a temperature in the range of 50-130°C, followed by a second stage at a temperature above the first stage and in the range of 70-155°C, and then optionally at a temperature above the The third stage is carried out at the temperature of the second stage temperature and at a temperature higher than or equal to 120°C, preferably 120-155°C. In this preferred process, the hydrogenation is optionally carried out in two or more separate reaction vessels.

Hydrogenation catalysts useful for the conversion of aldehyde functional groups (such as occur in HPA) or aldehydes formed by acetal inversion to alcohol functional groups (such as in PDO) include Group VIII metals such as nickel, cobalt, iron , palladium, rhodium, ruthenium and platinum. Copper is also known to be active.

A preferred such hydrogenation catalyst is described in US 5945570 (incorporated herein by reference). Preferred hydrogenation catalysts contain as main active component 25-60% by weight of nickel (in the form of Ni ⁰ ), preferably 30-45% by weight. Nickel within the active catalyst is predominantly in reduced form. The catalyst contains 5-20 wt% molybdenum (Mo ^O form), preferably 6-16 wt%. Molybdenum exists in the catalyst in the form of metal and oxide. Molybdenum has a binding function and is also an activity promoter.

The binder portion of the catalyst acts as a "binder", holding the discrete components together and providing resistance to crushing against pressure drop across the catalyst bed. The binder accounts for 30-70% by weight of the catalyst and consists of oxides and silicates of silicon and oxides of zinc, zirconium, calcium, magnesium and/or aluminum. Typically, the catalyst contains: 30-70 wt% silicon, preferably 35-55 wt%; 0-2 wt% zinc, preferably 0-1 wt%; and 0-2 wt% aluminium. The binder contains no more than 2 wt% calcium, preferably 0-1 wt% calcium. Preferred embodiments of the catalyst are substantially free of zinc or calcium. For the first stage hydrogenation of 3-hydroxypropanal to 1,3-propanediol in aqueous solution, the preferred catalyst composition contains 35 wt% nickel and 8-12 wt% molybdenum, with the balance of binder material as above stated.

The catalyst can be prepared by any procedure which combines the active nickel component, the molybdenum component and the binder material in solid bulk form. A preferred catalyst preparation involves mixing nickel oxide, binder material such as attapulgite and molybdenum trioxide powder into a homogeneous powder. Next, a solution of colloidal silica in sufficient water is stirred within the solid mixture to form an extrudable mixture. The wet mixture is then extruded through a die plate with a pore diameter of 0.040-0.070″. The extrudate is dried at 100-125° C. for a time sufficient to reduce the moisture content to less than 5 wt%. Then in air, at 450-550 The dried extrudate is calcined at 350°C for at least 3 hours until the desired strength is achieved. Before use, the catalyst is reduced under hydrogen at a temperature in the range of 350-450°C for a time sufficient to reduce at least 60% of the Nickel. If the reduced catalyst is not used immediately, it is cooled to ambient temperature and stored in a protective medium such as 1,3-propanediol until use.

In one embodiment of the invention, an acidic cocatalyst is added in the subsequent hydrogenation stage. The acidic co-catalyst facilitates the hydration reaction of the MW132 acetal back to PDO and HPA, thus increasing the yield of PDO in the overall hydrogenation process. Acidic promoters useful in the present invention may be selected from acidic zeolites, acidic cation exchange resins, heteropolyacids, amphoteric metal oxides exhibiting increased acidity relative to the hydrogenation catalyst used, and soluble acids. They are usually added in a ratio of 1:5 to 2:1 to the fixed bed or slurry hydrogenation catalyst. In a highly preferred embodiment, the acidic promoter is added in the third stage of hydrogenation and/or used in the post-hydrogenation stage of the initially hydrogenated aqueous PDO product.

It is also possible to include an acid promoter-promoted nickel hydrogenation catalyst in a fixed bed in all hydrogenation stages.

Alternatively, the acidic co-catalyst may be a soluble acid formed as a co-product or added during the process to achieve a pH of 2-5.5.

Preferred zeolite catalysts contain one or more modified zeolites, preferably acid modified zeolites. These zeolites should contain a pore size large enough to allow ingress of acyclic or aliphatic compounds. Preferred zeolites include, for example, zeolites of the MFI structure type (e.g. ZSM-5), MEL (e.g. ZSM-11), FER (e.g. Ferrierite and ZSM-35), FAU (e.g. zeolite Y), BEA (e.g. beta), MFS (e.g. ZSM-57), NES (e.g. NU-87), MOR (e.g. mordenite), CHA (e.g. chabazite), MTT (e.g. ZSM-23), MWW (e.g. MCM-22 and SSZ-25), EUO (e.g. EU-1, ZSM-50 and TPZ-3), OFF (e.g. zeolite), MTW (e.g. ZSM-12) and zeolites ITQ-1, ITQ-2, MCM-56, MCM-49, ZSM-48, SSZ-35, SSZ-39 and mixed crystalline phase zeolites such as zeolite PSH-3. Various zeolites can be found in "Atlas of Zeolite Structure Types" by M.W.Meier, D.H.Olson and Ch. Baerlocher, published by Butterworth-Heinemann (4th revised edition, 1996) (published on behalf of the Structure Committee of the International Zeolite Society) Synthesized structure types and references. Structure types and references of the above mentioned zeolites are available on the Internet at www.iza-structure.org. These zeolites are commercially available from Zeolyst International, Inc. and ExxonMobil Corporation. More preferably, the zeolite is a crystalline aluminosilicate which may contain traces of boron from the feedstock in the absence of deliberate addition of a boron source to enrich the boron content.

Other suitable catalysts include acidic cation exchange resins. These ion exchange resins include gel-type or macroreticular (macroporous) ion exchange resins having acidic sulfonic acid functional groups, where the sulfonic acid functional groups are directly or indirectly bonded to the organic polymer backbone. Examples include: AMBERLITE(R) or AMBERLYST(R) A200, A252, IR-118, IR120, A15, A35, A36, XN-1010 or uniform particle size A1200 resins from Rohmand Haas; Dow MSC-1 or DOWEX(R) 50 series resins; SYBRON(R) C-249, C-267, CFP-110 resins; PUROLITE(R) C-100 or C-150 resins; RESINTECH(R) CG8; IWT C-211; SACMP; Another example of these cation exchange resins are NAFION(R) acidified perfluorosulfonic acid polymers.

Water resistant solid acid catalysts were recently reviewed by T. Okuhara in Chemical Reviews 2002 , published by the American Chemical Society, pp.3461-3665. Heteropolyacids are formed from two or more different oxyanions, and typically include tetrahedrally coordinated silica or phosphorus bonded through hexavalent molybdenum, tungsten, uranium, niobium, or tantalum. Acidic anions with cerium (Cs2.5) show high activity for acid-catalyzed reactions in the presence of water.

Amphoteric or acidic oxides including alumina, silica, silica-alumina, titania, zirconia and their variants (including sulfated zirconia, zirconium tungstate, phosphate, molybdate and niobate) .

Soluble acid promoters may include mineral acids such as hydrochloric, hydrobromic, hydrofluoric, or hydroiodic acid, and nitric acid. Weak acids such as phosphoric acid, or carboxylic acids such as acetic acid and propionic acid can be used.

The metal for the fixed bed hydrogenation catalyst may be a Group VIII metal such as nickel, cobalt, ruthenium, platinum or palladium as well as copper, zinc, chromium and mixtures and alloys of these metals. A preferred catalyst is a particulate nickel-based composition on a water stable support such as a ceramic, such as the preferred catalysts described in US 5945570. The particle size of this catalyst is consistent with fixed bed operation, so typically ranges from 100 microns to 3 mm, with larger particles resulting in lower pressure drop at the expense of activity.

Example 1

Hydrogenation of PDO aqueous solution containing relatively high content (about 5 wt%) of MW132 acetal prepared with ethanol as a co-solvent (67 wt%) in a 500 ml batch autoclave reactor to evaluate the reversion of MW132 acetal and other similar impurities to PDO .

This catalyst (the preferred fixed-bed-promoted nickel hydrogenation catalyst described in US 5945570) showed no acetal reversal within 5 hours at 150 °C and 1000 psi _H2 , despite using 7.8 wt% catalyst (g-catalyst /g-liquid mixture). The reaction was terminated after a total of 11 hours and 1.5 wt% acidic USY zeolite was added. Over the next 2 hours, 75% of the MW132 acetal reversed to PDO despite a 20°C drop in temperature to 130°C. After an additional 2.5 hours at 150° C. with the acidic zeolite as cocatalyst, a final acetal conversion of 95% was obtained. This result clearly shows that the addition of a zeolite co-catalyst has a beneficial effect on the MW132 acetal reversal rate.

Example 2-17

A series of batch hydrogenation experiments were carried out under 1000 psi of hydrogen in a 500 ml autoclave using 240 g of a mixture of 2.7 wt% MW132 acetal in 31 wt% PDO and a rest solvent containing water (Table 1). In all cases a total of 4 g of catalyst was used. For "base case nickel" 4 g of the preferred catalyst described in US 5945570 was used as hydrogenation catalyst. In the case of co-catalyst, the amount of hydrogenation catalyst was reduced (2 g) and an equal amount of co-catalyst was used, resulting in a total of 4 g of combined catalyst. Periodic samples were taken for gas chromatographic analysis over a 3-5 hour period to monitor the reaction of the MW132 acetal and the formation of n-propanol (NPA), the major unwanted by-product of the reversion of the acetal to PDO.

Calculating the rate constant k as the slope of the conversion rate of MW132 versus time, for a given time "t" can be expressed as:

$\mathrm{<span\; class="notranslate">\; k</span>}<span\; class="notranslate">\; =</span>\frac{<span\; class="notranslate">\; -</span>\mathrm{<span\; class="notranslate">\; ln</span>}<span\; class="notranslate">\; (</span>\mathrm{<span\; class="notranslate">\; 1</span>}<span\; class="notranslate">\; -</span>\mathrm{<span\; class="notranslate">\; x</span>}<span\; class="notranslate">\; )</span>}{\mathrm{<span\; class="notranslate">\; t</span>}{<span\; class="notranslate">\; \&Center\; Dot;</span>\mathrm{<span\; class="notranslate">\; f</span>}}_{\mathrm{<span\; class="notranslate">\; cat</span>}}}$

X is the fractional conversion of MW132 acetal, or

$\mathrm{<span\; class="notranslate">\; x</span>}<span\; class="notranslate">\; =</span>\mathrm{<span\; class="notranslate">\; 1</span>}<span\; class="notranslate">\; -</span>\frac{<span\; class="notranslate">\; [</span>\mathrm{<span\; class="notranslate">\; MW</span>}\mathrm{<span\; class="notranslate">\; 132</span>}<span\; class="notranslate">\; ]</span>}{{<span\; class="notranslate">\; [</span>\mathrm{<span\; class="notranslate">\; MW</span>}\mathrm{<span\; class="notranslate">\; 132</span>}<span\; class="notranslate">\; ]</span>}_{\mathrm{<span\; class="notranslate">\; 0</span>}}}$

Where the subscript "0" indicates the initial concentration. The parameter f _cat represents the weight fraction of the catalyst or the weight of the total catalyst (typically 4 g) divided by the total mass of the liquid solution (typically 240 g).

The selectivity to undesired n-propanol is also defined as the mass of propanol formed relative to the mass of the reversed MW132 acetal (mainly relative to PDO and n-propanol).

Considering that an optimal high-temperature hydrogenation catalyst for good space-time yield will optimize the rate and selectivity of the desired product (PDO), the ratio of the reaction rate of MW132 divided by the (undesired) propanol selectivity is defined as the production rate factor "P". Higher values of P therefore reflect more economical performance for high temperature inversion of acetals.

Runs 2 and 3 of Table 1 represent reproducibility of the base Ni catalyst without the addition of co-catalyst. For runs 3-8 ("Series A"), an amphoteric or acidic oxide cocatalyst was added in place of an equivalent amount of the base nickel hydrogenation catalyst. The results show that for the oxides zirconia, titania, silica, alumina and silica-alumina, the use of a co-catalyst increases the rate of reversal of the MW132 acetal. For run #9 ("A series"), the addition of a strongly acidic NiO/ _W03 /alumina cocatalyst also increased the rate of acetal reversal relative to the base composition without the addition of an acid cocatalyst.

By the back titration method, an aqueous solution of 0.1N potassium hydroxide or potassium carbonate was contacted with a solid acid cocatalyst overnight at room temperature (while mixing), and then the supernatant was titrated with 0.1N HCl to evaluate the amount of residual KOH, thereby evaluating Acidity of several solid oxide promoters in water. The results of this method are shown in Table 2. Both the oxide-promoted and NiO/ _WO3 /alumina catalysts tested were significantly more acidic than the nickel-based catalysts, explaining their increased activity in the reversal of MW132. For these experiments, the amount of base (potassium hydroxide or potassium carbonate) was limited to avoid dissolution of the solid oxides. HP Boehm, "Chemical Identification of Surface Groups" in Advances in Catalysis, Vol. 16, pp. 179-273 (1968) describes back titration of solid oxides in aqueous media for the determination of acidity.

In series "A" of Table 1, an increase in the formation of n-propanol is evident upon addition of the acid co-catalyst. The "B" series examines the case where the hydrogenation temperature is reduced to 130°C, which results in a decrease in the amount of undesired n-propanol formed, while still providing increased M.W.132 Acetal Reversal Speed. The possibility of temperature drop is particularly attractive since it is known that the stability of fixed bed hydrogenation catalysts is negatively affected by temperature in an aqueous environment. For the same space-time conversion ratio, the present invention can achieve improved speed at lower temperature than required, which in turn can prolong the life of high temperature hydrogenation catalyst.

Series C and D (Table 1 ) examine the case of a soluble organic acid (3-hydroxypropionic acid) as an acidic promoter for Ru/alumina catalysts and basic Ni catalysts. For both cocatalysts, the reverse reaction of MW132 acetal was accelerated when the reaction mixture was made more acidic by pH adjustment. This result indicates that the conversion of acetal is a function of proton (H ⁺ ) concentration, which is supported by pH. A pH tolerant catalyst in combination with low pH (2-5) in the high temperature hydrogenation stage results in increased back conversion of the acetal heavy fraction. The operating pH range for Ni catalysts is 4-7.5, but is preferably kept above 5.0, preferably 5.2-5.5, since otherwise the lifetime of the catalyst is negatively affected. Operating in the 4-5 range yields a greater increase in reverse speed.

Series E examines a Raney cobalt slurry catalyst (W.R. Grace #18659-21) operated at a lower temperature of 120°C, which is suitable for use with a sulfonic acid resin (sulfonated polystyrene Amberlyst(R) 15 from Rohm and Haas ) as a cocatalyst. No reversal of the M.W. 132 acetal was measured at this temperature in the absence of a strong acid resin cocatalyst. With the resin cocatalyst, despite operating at a temperature 20°C lower than the base case without the acidic cocatalyst, the rate was more than twice that of the nickel base and only 1/2 the undesirable Amount of n-propanol. Lower temperatures will help prolong catalyst life while maintaining the ability to reverse acetalization to PDO.

Table 1: Examples 2-17: Acetal Reversal Reactions Using Acidic Cocatalysts series Ex. hydrogenation catalyst Co-catalyst Hydrogenation catalyst g Acid cocatalyst g temperature °C pH NPA selectivity % MW132 Speed (1/h) P=speed/npa1/h/% AAAAAAAAA 23456789 Ni basic ingredients Ni basic ingredients Ni basic ingredients Ni basic ingredients Ni basic ingredients Ni basic ingredients Ni basic ingredients Ni basic ingredients Zirconia Free Titania 55/45 Silica/Alumina Alumina Silica NiO/WO ₃ /Alumina 44222222 00222222 140140140140140140140140 5.55.55.55.55.55.55.55.5 11.710.814.216.719.622.524.222.3 12.17.324.244.245.034.440.854.8 1.00.71.72.62.31.51.72.5 BBBB 231011 Ni Basic Components Ni Basic Components Ni Basic Components Ni Basic Components None None None 55/45 Silica/Alumina 55/45 Silica/Alumina 4422 0022 140140140130 5.55.55.55.5 11.710.819.65.2 12.17.345.019.8 1.00.72.33.8 CC 1213 Ru/alumina Ru/alumina C3-Acid No 44 no no no 140140 3.95.5 5.66.9 159.726.5 28.73.8 DDDD 141523 Ni Basic Components Ni Basic Components Ni Basic Components Ni Basic Components C3-Acid C3-Acid None None 4444 0000 140140140140 4.455.55.5 10.712.711.710.8 42.332.712.17.3 3.92.61.00.7 EE 1617 Raney Co Raney Co A15 resin no twenty two 20 12120 5.55.5 5.90.0 23.20.0 4.00.0

Titanium Oxide: Norton #XT25376, 1/16"

Zirconia: Engelhard #803A-6-2-31 1/16 inch

Alumina: AX-200, CRI-Catalysts, 1/32 in.

Silica: Dayisil Grade 57, Grace Davison 60/80 mesh

Silica-Alumina: Ghent 654/CB01-149 CRI-Catalysts, 1/16 inch

A15: Amberlyst 15 Strong Acid Resin, Rohm and Haas

1/h=1/hour, airspeed unit

Table 2: Back titration of catalyst/cocatalyst acidity Catalyst / Cocatalyst Titrant Catalyst acidity meq/g Ni Basic Composition 55/45 Silica-Alumina Alumina (Ax200) Titanium Oxide NiO/WO ₃ /Alumina 0.1N KOH0.1N KOH0.1N KOH0.1N KOH0.1N KOH 0.680.930.950.860.98 Ni basic composition None 0.1N K ₂ CO ₃ 0.1N K ₂ CO ₃ 0.781.00

1.0meq/g=100% consumption of alkali titrant

Example 18

The binary continuous flow reactor was operated at 150° C. and a hydrogen pressure of 1180 psi (8135 kPa) with 2.0 g of a different batch of catalyst used in the previous examples. Set the liquid flow rate to obtain an average weight hourly space velocity (WHSV) of 0.9-1.3:

WHSV = g liquid feedstock/g catalyst/hour

Hydrogen was co-fed to the reactor inlet at a flow rate of 8-10 ml/min (measured by an outlet rotameter) at standard temperature and pressure, thereby providing operation in a three-phase state (gas-liquid-solid), This is expected for a fixed bed bubble column reactor. Samples were taken periodically for gas chromatography (GC) analysis of the MW132 acetal content of the effluent. The liquid feed solution consisted of 2.2-2.3 wt% MW132 acetal in a mixture of 27-31 wt% PDO in water, for reactor "A" the pH was adjusted to 4.9 by addition of 1 N KOH, and for reactor "B ", adjust the pH to 5.5. This example simulates the feedstock for the final high temperature hydrogenation stage. Rate constants for acetal reversal were calculated as described above. For a fixed bed catalyst, such as that used in this example, f _cat =1.

Table 3 shows the results of the 1-month test. Reactor A with a pH 4.9 feedstock showed less MW132 in the effluent than Reactor B operated with a pH 5.5 feedstock. The rate constants (calculated as described above) indicated that the catalyst activity at pH 4.9 was about twice that observed with the feedstock at pH 5.5. This experiment demonstrates that the advantage of operating at lower pH to increase the speed of cyclic acetal reversal is indeed maintained during industrial long-term continuous operation.

Table 3: Effect of pH on acetal reversal in continuous flow microreactor experiments Reactor A: Feedstock 2.22 wt% MW132 at pH 4.9 Reactor B: Feedstock 2.29 wt% MW132 at pH 5.5 sky Total raw material g/g catalyst WHSV Export MW132wt% Rate constant 1/h Total raw material g/g catalyst WHSV Export MW132wt% Rate constant 1/h 0.006.007.008.009.0012.0026.0027.0028.0029.00 0.0137.1159.4182.6206.2267.4649.0676.2703.5731.1 None 0.950.930.970.980.851.141.131.141.15 None 0.2950.1660.1120.0550.0460.0010.0010.0010.001 None 1.922.412.893.643.298.758.738.788.85 0.0157.6177.4192.1219.2270.5795.2818.9841.4871.6 None 1.090.830.611.130.711.560.990.941.26 None 0.4380.2650.1240.2510.1050.2480.1250.1310.154 None 1.781.751.772.462.173.422.842.653.36

Claims (17)

1. prepare the method for 1,3-propanediol by hydrogenation of 3-hydroxy propionaldehyde, the method may further comprise the steps: (a) preparing an aqueous mixture of 3-hydroxypropanal, (b) passing the aqueous 3-hydroxypropanal mixture through a hydrogenation zone, wherein the hydrogenation zone consists of at least two stages, in the range of 50-130°C in the presence of a fixed bed or slurry hydrogenation catalyst The first stage of hydrogenation is carried out under hydrogenation conditions at a temperature higher than that of the first hydrogenation stage and in the range of 70-155° C., with the addition or presence of an acidic cocatalyst in at least one subsequent stage The subsequent stage of hydrogenation is carried out under hydrogenation conditions to form an aqueous solution of 1,3-propanediol, wherein the acidic co-catalyst is selected from acidic zeolites, acidic cation exchange resins, acidic or amphoteric metal oxides , heteropolyacids and soluble acids; and (c) recovering said 1,3-propanediol. 2. The process of claim 1, wherein the hydrogenation zone consists of at least three stages, wherein the second stage is carried out under hydrogenation conditions at a temperature higher than the temperature of the first stage in the range of 70-155°C Hydrogen, and the addition or presence of an acidic co-catalyst in the final stage of the hydrogenation zone, wherein the hydrogenation of the final stage is carried out under hydrogenation conditions at a temperature higher than that of the second stage and at a temperature greater than or equal to 120°C. 3. The method of claim 2, wherein the temperature of the second stage is 70-140°C, and the temperature of the last stage is 120-155°C. 4. The method of any one of claims 1-3, wherein the acidic promoter is selected from the group consisting of acidic zeolites, acidic cation exchange resins, acidic or amphoteric solid metal oxides, heteropolyacids and soluble acids. 5. The method of any one of claims 1-4, wherein the pH of step (b) is higher than 5. 6. The process of any one of claims 1-5, wherein a slurry hydrogenation catalyst and an acidic cation exchange resin are used, and the temperature of the final hydrogenation stage is 100-140°C. 7. The method of any one of claims 1-6, wherein the aqueous 3-hydroxypropionaldehyde mixture passed to the hydrogenation zone is an aqueous solution containing 3-hydroxypropionaldehyde, wherein the concentration of 3-hydroxypropionaldehyde is based on the weight of the aqueous solvent It is 3-50wt%. 8. The method of claim 7, wherein the aqueous 3-hydroxypropanal mixture passed to the hydrogenation zone is an aqueous solution containing 3-hydroxypropanal, wherein the concentration of 3-hydroxypropanal is 10-40wt based on the weight of the aqueous solvent %. 9. The method of claim 7, wherein the hydrogenation is carried out using a fixed bed catalyst, and the concentration of 3-hydroxypropanal in the aqueous 3-hydroxypropanal mixture passed to the hydrogenation zone is 0.2-15 wt% based on the weight of the aqueous solvent. 10. The method of claim 9, wherein the aqueous 3-hydroxypropanal mixture of step a) is diluted to 0.2-15 wt%, based on the weight of the aqueous solvent, by adding an aqueous solution of 1,3-propanediol. 11. The method of claim 10, wherein the total amount of 3-hydroxypropionaldehyde and 1,3-propanediol in the diluted mixture is 20-40 wt%, based on the weight of the aqueous solvent. 12. The method of claim 11, wherein the aqueous solution of 1,3-propanediol is cooled before being added to the aqueous 3-hydroxypropanal mixture of step a). 13. The method of claim 9, wherein the hydrogenation is carried out using a fixed bed catalyst, and the concentration of 3-hydroxypropanal in the aqueous 3-hydroxypropanal mixture passed to the hydrogenation zone is 0.5-8 wt% based on the weight of the aqueous solvent. 14. The method of claim 13, wherein the aqueous 3-hydroxypropanal mixture of step a) is diluted to 0.5-8 wt%, based on the weight of the aqueous solvent, by adding an aqueous solution of 1,3-propanediol. 15. The method of claim 14, wherein the total amount of 3-hydroxypropionaldehyde and 1,3-propanediol in the diluted mixture is 20-40 wt%, based on the weight of the aqueous solvent. 16. The method of claim 15, wherein the aqueous solution of 1,3-propanediol is cooled before being added to the aqueous 3-hydroxypropanal mixture of step a). 17. The process of claim 1, wherein a fixed bed hydrogenation catalyst is used in all hydrogenation stages and an acidic cocatalyst is added together with the hydrogenation catalyst.